The dehydrogenation of light hydrocarbons can be performed several ways. In the following, the dehydrogenation of propane will be examined, however, the principles described herein apply equally to the dehydrogenation of other light hydrocarbons. Light hydrocarbons are defined as having five carbon atoms or less, i.e., ethane, propane, butane, isobutane, pentane and isopentane.
It is known in the art to dehydrogenate propane to propylene such as by endothermic vapor phase dehydrogenation. The dehydrogenation can occur either thermally or catalytically. Simultaneously, other thermal reactions occur which crack and/or degrade the propane molecule and produce compounds smaller than propylene. Also, propylene itself is subject to unwanted thermal degradation reactions.
The principal reactions in propane dehydrogenation include the following: ##STR1## All of these reactions require substantial quantities of heat per pound of feed.
Propylene is produced commercially by the thermal cracking of propane in pyrolysis coil cracking reactors, along with substantial quantities of ethylene and methane; by the catalytic dehydrogenation of propane; and by the recovery of co-produced propylene from the effluent when cracking heavier compounds either thermally or catalytically. Catalytic dehydrogenation typically occurs in fixed bed reactors where propane is passed over a catalyst bed and reaction occurs.
The choice of process depends upon the selectivity and/or the conversion of propane to propylene desired. Thermal cracking reactions are not selective to propylene and the process yields substantial quantities of co-produced ethylene and methane. Thermal cracking can however achieve complete conversion of the feedstock thus eliminating recycle of the feedstock.
If propylene selectivity is desired, i.e., higher yields of propylene per unit propane feed, catalytic dehydrogenation is preferred. The catalysts allow the dehydrogenation reaction to occur at substantially lower temperatures than required for thermal dehydrogenation. The lower temperatures suppress unwanted thermal cracking reactions and high yields of propylene per unit propane are obtained. At these lower temperatures, however, absolute conversion is limited by equilibrium. Thus, while selectivity to propylene is high, conversion is low. Such a process thus requires substantial recycle of unreacted feed.
Several commercial processes exist for catalytically dehydrogenating hydrocarbons using fixed bed technologies. Such processes are carried out either adiabatically or isothermally.
There are two major commercial processes using adiabatic conversion over a catalyst. In one of the processes, the heat for the endothermic dehydrogenation reaction is provided by preheating the reactor feed to temperatures greater than the reactor outlet. The propane feed is highly diluted with an inert diluent (steam or hydrogen/methane). This diluent acts primarily as a heat carrier but also serves to minimize catalyst fouling. Catalyst is gradually removed from the bed, regenerated, and returned. The rate of circulation is quite low; bed turnover is measured in days and is not a source of heat. This "moving bed" operation is quite complex mechanically and limited to very low circulations.
Since the dehydrogenation reaction requires substantial quantities of heat per unit propane, failure to dilute the mix would require the propane to be preheated to a very high temperature, causing extensive thermal reactions prior to entering the catalyst bed. This would result in a system with poor selectivity to propylene. To compensate, the flow of diluent is typically equal to or greater than the flow of feed. High diluent flows reduce the capacity of the processing equipment.
The dehydrogenation is normally carried out in two or three stages. The partially reacted feed mixture is reheated in each stage to provide the heat necessary to complete the reaction. Interstage temperatures of between 565.degree. C. and 630.degree. C. are necessary to obtain commercially acceptable conversion levels and to minimize thermal cracking of the feed.
Even with high dilution levels, equilibrium limits the reaction to only about 40% conversion per pass at temperatures required for high selectivity. This requires that the unreacted material be recycled if the goal is to completely convert the fresh feed. At 40% conversion per pass, 1.5 lb of recycle is fed to the reactor for every 1.0 lb of fresh feed. Low conversion and high dilution represent an inefficient use of equipment capacity since effectively only 40% of the reactor is used without even considering the dilution. With a one to one molar dilution rate, only 20% of the potential reactor capacity for fresh feed is used. Furthermore, the high recycle rates and high dilution require extensive downstream separation equipment.
Another adiabatic conversion process passes the feed mixture over a highly preheated fixed catalyst bed. As the reaction proceeds, the catalyst gradually gives up its heat to reaction and cools, leading to changes in conversion over time. Eventually (time measured in minutes), the extent of conversion becomes too low for reasons of control and the reactor is taken off line. A second reactor, which was being preheated by direct contact with hot combustion gases while off-line, is then switched on-stream. A continuous process is achieved by the cycling of several reactors simultaneously. This involves complicated valving and control systems. Conversions in this process are slightly higher (on average 60%) than the previous process but the process still requires substantial recycle. Further, conversions change with time and products of combustion (from the reheat cycle) contaminate the product gas flow.
In a non-adiabatic commercial process the catalytic dehydrogenation of hydrocarbons is carried out in a gas fired tubular reactor. Conversions of only about 40 mole percent are generally achieved in such reactors. In addition, this process requires the recirculation of large quantities of hot gas thereby consuming substantial amounts of energy. Also, many tubes are required to achieve high capacities in this process.
Another non-adiabatic process involves the catalytic fixed bed dehydrogenation of hydrocarbons in a nearly isothermal fixed bed reactor. (See, e.g., U.S. Pat. No. 4,287,375, Moller, et al.) The tubular reactor is immersed in an isothermal molten salt bath maintained at a temperature of 600.degree. C. while the reaction temperature is about 20.degree. C. less than the molten salt bath. In this process, a water vapor/ethylbenzene mixture at a ratio of about 1.2-1.5 kg steam/1.0 kg hydrocarbon is introduced into the dehydrogenation reactor. High dilution levels result in low feedstock capacity.
Both of these non-adiabatic processes suffer from difficulties in providing heat to the fixed bed of catalysts. One method of adding heat is to preheat the feed and/or diluent as mentioned above. This technique has a negative effect on reactor capacity and selectivity. The second method is to divide the bed into a plurality of smaller beds, i.e. tubes, and allow heat to flow through the walls of the tubes. This requires complex mechanical systems, but has been achieved as referenced above.
It is even more difficult to achieve short kinetic residence times in these fixed bed systems. As gas flows over a bed of catalyst, pressure is lost. In order to avoid large unfavorable pressure drops, the gas velocity must be low and catalyst particle sizes must be large. With low gas velocity, the fixed bed must be shallow yet have considerable cross-sectional area to retain high capacity. This leads to distribution problems not only for feed input into the catalyst bed but for distribution of product cooling media to stop the unwanted thermal side reactions from continuing once the product has left the catalyst bed. For the tubular reactors mentioned above, distribution is especially critical. Also, the large catalyst particles required for low pressure drop create significant diffusional resistances to reaction which precludes operation at short residence times.
There is a further consideration of residence time since all reactions continue until either the feedstock is exhausted or the process operating conditions have changed. Catalytic reactions are usually stopped by having the product gas exit the catalyst bed, however, this does not stop the thermal reactions. Heat must be removed from the system or a quench medium added to the product gas in order to reduce the temperature and stop the thermal reaction. Fixed bed systems, with either their plurality of tubes or high cross-sectional area, are not well suited for rapid heat removal or rapid introduction of a quench medium to cool the product gasses and prevent further reaction.
Low kinetic residence times are therefore difficult to achieve in conventional dehydrogenation reaction systems and, as such, these systems cannot reach a favored reaction temperature without a loss of selectivity. These conventional systems also have limitations involving heat input, feed distribution and pressure drop regardless of the feed.
It would therefore represent a notable advance in the state of the art if a dehydrogenation process could be provided which operates at high temperature, short kinetic residence time, low pressure drop, low feed dilution levels, high conversion and high selectivity, and which can facilitate catalyst regeneration.